Hydrocracking of distillates



Dec. 3, 1968 P. c. KEITH ET AL 3,414,505

HYDROCRACKING OF DISTILLATES Original Filed Nov. 21, 1963 HYDROGEN 36 34 H RECYCLE 55 GAS Q a 30 32 1F 54 GAS 5| H2 SEPARATOR 50 20 57 LT. NAPHTHA RED I V 60 a as REACTOR HVY. NAPHTHA E GAS z SEPARATOR g 2 LT. 0| T. 4| 5 S 62 :72

HVY. DIST.

FEED V O lo l2 J 22 RECYCLE OIL INVENTOR.

ATTORNEY Patented Dec. 3, 1968 3,414,505 HYDROCRACKING F DISTILLATES Percival 'C. Keith, Peapack, and Michael C. Chervenak,

Pennington, N.J., assignors to Hydrocarbon Research,

Inc., New York, N.Y., a corporation of New Jersey Continuation of application Ser. No. 325,300, Nov. 21,

1963. This application Nov. 9, 1966, Ser. No. 593,037

2 Claims. (Cl. 208-108) ABSTRACT OF THE DISCLOSURE This invention pertains to a highly flexible distillate cracking process for making a wide array of distillate and gasoline products of completely satisfactory quality. It comprises essentially the conversion of liquid gas oil having a boiling range of about 325 to about 1100 F. to a high quality heater oil and desulfurized lighter material by use of an ebullated bed technique employing a dual function catalyst.

This invention, which is a continuation of the invention shown in our application, Ser. No. 325,300, filed Nov. 21, 1963, now abandoned, relates to hydrogen conversion of hydrocarbon distillates; particularly it relates to hydrocracking of petroleum distillates to gas oils or to heavierthan-gasoline light distillate.

The petroleum refiners have struggled for years with the seasonal nature of gasoline and domestic heater oil demands, demands which peak in opposed parts of the year. Improvement in refining operations have alleviated to some extent this problem but it is still necessary to store gasoline made in the winter and spring to meet the summer demand; the winter demand of heater oil is met by storing oil produced in the summer and fall. An enormous amount of capital is tied up in storage tanks and in the value of stored product, because the presently known distillate cracking processes are not flexible enough.

Superimposed on the above is a demand for diesel oil and an ever increasing demand for jet fuel, especially jet fuel boiling in the kerosene and diesel oil range. Present processes are aimed at gasoline as the desired product; the octane race has resulted in the use of processes which produce large percentage of aromatic hydrocarbons because these have very high octane numbers. Simultaneously the aromaticity of the higher boiling products is increased.

Good quality kerosenes require a high saturated hydrocarbon content because aromatic hydrocarbons burn with a smoky flame. Good quality diesel oils require a high cetane number which is only given by saturated hydrocarbons; aromatic hydrocarbons also increase carbon formation in the engine. Jet engines require saturated hydrocarbons to keep combustion carbon deposits down and to give more B.t.u.s per pound of fuel.

The presence of organo-sulfur compounds in gasoline is deleterious because of odor and the decrease in effectiveness of tetra alkyl lead anti-knock agents. Even where odor is not a major consideration, it is now known sulfur compounds are undesirable because of their corrosiveness and adverse effect on stability of the products.

While the presently known hydrodesulfurization processes can overcome the sulfur problem and the presently known hydrocracking processes can impart some greater flexibility of refinery operations, these processes still have drawbacks either in flexibility, or in operating effectiveness, or in economic factors.

An object of this invention is a hydrocracking process for conversion of hydrocarbon distillates, particularly heavy gas oils.

Still another object of this invention is a distillate hydrocracking process which accomplishes the cracking while simultaneously decreasing the organo-sulfur and organo-nitrogen content thereof.

Other objects of the invention will become apparent in the course of the detailed description thereof.

The figure shows a schematic flow sheet of a process for hydrocracking gas oil to essentially high quality heater oil and lighter material.

The invention is described in connection with the drawing which forms a part of this specification. Only those items of equipment which are necessary to an understanding of the invention have been included in this drawing. The pumps, valves, heat exchangers, fractionators, etc. needed in practice can be easily added by anyone-skilled in this art.

Briefly this process comprises: introducing essentially liquid gas oil boiling in the range of about 325 F. to about 1100 F. to the lower portion of a catalytic reaction zone; passing said liquid gas oil and hydrogen gas upwardly through a body of solid particulate catalyst, said catalyst being characterized by the ability to promote at least hydrogenation reactions, and preferably the dualfunction ability to promote both hydrogenation reactions and cracking reactions; said body being characterized by a fairly, well-defined upper level of solid particles; said hydrogen gas being charged in an amount at least that needed to obtain the desired degree of conversion and product quality; said catalytic reaction zone being maintained at a temperature of between about 500 F. and about 900 F.; withdrawing a conversion product mixture from said reaction zone at a point above said upper level of solid particles; distilling said conversion product mixture to separate the material boiling below about 750 F. from the heavier material; recycling said heavier material to said reaction zone as part of the liquid gas oil charge; and continuing said recycle operation until essentially all of the introduced gas oil has been converted to products boiling below about 750 F.

Feed hydrocarbon oil from source 10 is introduced by way of lines 12, 14 and 16 into the bottom of catalytic reactor 18. The feed oil to this process may boil in the range of about 325 F. to about 1100 F. This process is especially suitable for heavy gas oils which boil in the range of about 600 F. and about 1100 F. and include a major amount of material boiling above about 750 F.

The feed may be a distillate obtained from the distillation of crude oil, a reduced crude, or a wide boiling range distillate. Also the feed may be a product of a conversion process, for example, thermal cracking of gas oil; visbreaking of reduced crude; coking of heavy distillates, reduced crude and other residual material, such as cracked tar; catalytic cracking of distillates-fixed bed, moving bed or fluid bed; hydrogen treating and cracking processes whether applied to distillates or reduced crudes or other residual materials. The feed may be a mixture of two or more of the aforesaid distillates.

Hydrogen gas is passed by way of line 20 into line 14 and along with the feed from source 10 and recycle oil from line 22 is introduced into a lower portion of reactor 18, below the lower level, formed by distributor 24, of catalyst bed (body) 26. The hydrogen gas may be pure, or a recycle gas containing appreciable amounts of methane, nitrogen, hydrogen sulfide, etc. It is preferred that the hydrogen sulfide be as low as economies and operations permit. Since the conversion reaction is in part controlled by hydrogen partial pressure, it is desirable to keep the diluent gas content of the hydrogen stream as low as practical in order to decrease the pressure requirements in reactor 18. Here outside hydrogen from source 30 is passed by line 32 to line 20; recycle hydrogen from compressor 34 and line 36 is passed into line 20.

Sufiicient hydrogen is charged to provide the amount needed to obtain the desired degree of conversion and product quality. In general, the hydrogen consumption in this process is about 700-2500 standard cubic feet per 42 gallon barrel (s.c.f./b.) of feed introduced from source 10. More than the consumption amount of hydrogen will be passed through reactor 18; the circulation amount of hydrogen may range from about 3000 to about 15,000 s.c.f./b. of introduced feed.

The lower limit of the catalyst body 26 in reactor 18 is determined by distributor 24. Herein this is shown as a perforator plate or grid of conventional type. Distributor 24 may be any of the types used to permit introduction of a liquid-gas feed below a body of catalyst.

A body of solid particulate catalyst 26 is maintained in reactor 18. This body, during operation of the process is placed in an ebullated state (also spoken of as an ebullated bed) with the particles in random motion in the liquid. The use of an ebullated bed of solid particles in gas-liquid contacting is described in detail in Johanson U.S. Patent No. 2,987,465, granted June 6, 1961 (Reissue No. 25,770). Briefly, in an ebullated bed the solid particles occupy a greater volume of space than in their settled state; they are in random motion as a result of the action of the upflowing stream of liquid and gas; and the bed is characterized by a fairly well-defined upper level of solid particles, i.e., there is no substantial carryover of solid particles by the liquid withdrawn from the upper portion of the reactor.

in this embodiment, an ebullated bed of solid particulate catalyst is present in reactor 18; it is desirable that the ebullated bed have an expanded volume at least about greater than the settled state volume of the body. Generally the expansion is about l0l00% greater than the settled state volume. The upper level of the ebullated bed in reactor 18 is shown by the doubledashed line 40.

The size of the catalyst particles will be determined by the type of catalyst, the conditions of operation in reactor 18, and by the degree of bed expansion desired. For an ebullated bed the particles may range from onehalf inch downward. Catalyst having sieve sizes of about 200 mesh have been used successfully. A particularly suitable particle size is afforded by a commercial catalyst which is available as cylinders: one-tenth inch diameter by onequarter inch long; one-sixteenth inch diameter by one-quarter inch long.

The catalyst used in this process is characterized by the ability to promote hydrogenation reactions and preferably the dual-function ability to promote both hydrogenation reactions and cracking reactions. (This type of catalyst has also been spoken of as acidic catalyst.) An especially suitable dual function catalyst is nickel-tungsten on a silica-alumina support. Cobalt-molybdenum on a silica alumina support is suitable. Numerous catalysts are known to have this dual-function characteristic and any of these may be used in this process. It is to be understood not all of these catalysts are equal and operating conditions will have to be adjusted to obtain optimum results with each catalystwithin the hereinafter defined operating conditions.

The catalyst also might be mono-functional in that it promotes substantially only hydrogenation reactions. Illustrative of these are metals such as nickel, cobalt, molybdenum, platinum, palladium on an alumina or silica gel support. A particularly suitable member of this class is cobalt molybdate on alumina. The dual-functional catalysts are desirable with the lower boiling range distillates; whereas the mono-functional catalysts work well with the higher boiling range distillates, such as heavy gas oil.

The upwardly flowing gas oil in reactor 18 is in the liquid state and while there may be as much as 90% vaporized, there is always sufficient oil in liquid state to maintain a continuous liquid phase. In general, sufficient pressure is maintained in the reaction zone to provide a hydrogen partial pressure in the order of about 1000 3000 psi. although still higher pressures may be used.

The catalytic reaction zone (reactor 18) is maintained at a temperature between about 500 F. and about 900 F., and particularly with dual functional catalysts at about 650 F.725 F. When the catalyst is mono-functional it is preferred to operate at a temperature of about 750-850 F.

It is desirable to have substantially an isothermal reaction zone. This is accomplished in part by the ebullated bed and by operating with recycle of oil to be retreated for additional conversion. In addition to the recycle of oil by way of line 22, oil may be withdrawn at 41 from an upper portion 42 of the reactor 18 above the ebullated bed 26 and circulated by pump 43 to the lower portion of the reactor below the catalyst bed 26.

This circulating stream is shown external to the reactor 18 or it may be an internal circulation through a conduit extending from below the catalyst bed to a point within zone 42 as described in the Johanson Reissue Patent, 25,770, and the internal circulation may be gravity flow or it may be assisted by an internal pump. The volume of internal oil circulation is dependent on the particular feed, catalyst and operating conditions and may vary from essentially zero to 20 or more volumes per volume of feed.

The upfiowing liquid oil is passed through the catalyst bed at the rate needed to obtain the desired catalyst ex pansion at the particular set of operating conditions. Broadly, the space velocity in the reactor 18 is between about 0.2 and 2 and commonly is about 0.7-1.5. Herein space velocity is defined as the volume of feed from source 10, at 60 F, charged per hour per volume of catalystin the expanded state in the case of an ebullated bed.

A great advantage of the ebullated bed lies in the ability to withdraw catalyst from bed 26 and add fresh or regenerated catalyst thereto while the operation is on stream, thereby maintaining more-or-less constant catalyst activity and consequent uniformity of product distribution over the length of the run. Such withdrawal and addition means are shown diagrammatically at 27 and 29. Replacement should be continuous to maintain a high level of activity and is usually in the order of about /2 to 5 percent (weight) per day.

It is to be understood that while only one reactor has been shown, two or more reactors may be operated in parallel or in series flow.

A mixture of conversion products and hydrogen gas is withdrawn from reactor 18 at a point above the upper level 40 of solid particles (zone 42 in reactor 18) and the mixture is passed by way of line 50 preferably being cooled in heat exchanger 51 to hydrogen gas separator 52. Here hydrogen gas, hydrogen sulfide by-product, and some non-condensable hydrocarbons are taken overhead by way of line 54 with part recycled to reactor 18 by way of compressor 34 and lines 36, 20, 14 and 16. A purge line, if needed, is shown at 55.

The recycle hydrogen gas stream may be treated to remove hydrogen sulfide and/or light hydrocarbons before being recycled to reactor 18. This treatment may be accomplished by any of the known processes.

From separator 52 the conversion product mixture is passed by way of line 56 through reduction valve 57 to gas separator 58. Here non-condensable gases such as methane, ethane and some propane are separated overhead and sent to disposal by way of line 60.

The liquid bottoms from separator 58 or alternatively, all of the bottoms from separator 52 if the gas separator 58 is not used, is passed by way of line 62 into fractionation zone 64. Zone 64 is shown as producing through line 66 a gas stream of methane, ethane, propane and some butane; an unstabilized light naphtha product through line 68 and a heavy naphtha product through line 70. The two naphtha streams represent the gasoline product of the process, which gasoline may have an end point as high as about 420 F.

In this process the liquid products boil below about 750 F. In this range lie kerosenes, jet fuels, diesel oil and the domestic heater oils and burner oils. Purely for illustration, zone 64 is shown as producing two distillate fractions boiling above the gasoline range and having an end point below about 750 F. A light distillate suitable for jet-fuel may be withdrawn by way of valved line 72; a heavy distillate suitable for burner oil use may be withdrawn by way of valved line 74.

It is to be understood that the cutpoint between the product stream and the recycle oil stream will be determined by the particular array of gasoline and heavier products which is desired. For example, a typical kerosenediesel oil medium jet fuel product has an end point of about 525 -550 F.; a heaver diesel oil or heater oil has an end point of about 625 650 F.; a No. 2 fuel oil has an end point of about 725 -750 F.

In this process, recycle oil boiling above the highest boiling desired product fraction is withdrawn from zone 64 and recycled by Way of lines 22 and 16 to reactor 18. The recycling operation is continued until essentially all of feed (introduced gas oil) has been converted to products boiling below about 750 F., or the end point of the highest boiling desired product whose end point is below 750 F.

Illustration I The process is illustrated by the hydrocracking of a heavy gas oil obtained as a byproduct of delayed coking of reduced crude. This 17 API coker gas oil boiling over the range of 640 F. to about 1050 B; it had a sulfur content of 2.9%.

The heavy gas oil feed is charged to reactor 18 along with recycle oilabout 0.34 volume per volume of feed and hydrogen gas by way of line 16. Reactor 18 is loaded with commercial cobalt-molybdate on alumina catalyst in the form of pills about A; inch diameter by inch thick. The upfiowing stream is controlled to provide an ebullated bed having about 20% expansion. The hydrogen gas flow is about 7000 s.c.f./ b. with a hydrogen consumption of about 2000 s.c.f./b. A circulation of oil is maintained in the reactor of about volumes per volume of heavy gas oil introduced. An essentially isothermal temperature of about 800 F., at a total pressure of about 1100 p.s.i., is maintained in catalyst bed 26.

The mixture withdrawn from upper zone 42 of reactor 18 has the hydrogen gas removed in separator 52 and the non-condensible gas removed in separator 58. The liquid product mixture is fractionated in a zone 64 into: a gas, propane and butane fraction; a light naphtha fraction; a heavy naphtha fraction-400 F. end point; a wide range distillate having an end point of about 650 F.; and a higher boiling recycle oil.

The ultimate yield of products boiling below 650 F. obtained by hydrocracking this coker gas oil to extinction is: methane and ethane, 1.6 weight percent; propane, 1.8 weight percent; butanes, 3.6 weight percent; light naphtha, 12.0 volume percent; heavy naphtha, 27.1 volume percent; and 400-650 F. distillate, 70.0 volume percent. The gasoline product has less than 0.1% sulfur and the heavier distillate is of satisfactory sulfur content and is suitable for use in kerosene, diesel oil or jet fuel. (It is also a superior feed to a catalytic cracking operation.)

The above description and illustration establishes that an extremely flexible distillate cracking process has been discovered suitable for making a wide array of distillate and gasoline products of completely satisfactory quality, at the choice of the refiner.

Thus having described the invention, what is claimed is: 1. A process for converting a distillate gas oil to essentially only products boiling below about 750 P. which process comprises:

introducing essentially liquid gas oil boiling in the range of about 600 F. to about 1100 F. and having a major amount of material boiling above about 750 F. to the lower portion of a catalytic reaction zone;

passing said liquid gas oil and hydrogen gas upwardly through a body of solid particulate catalyst at a velocity to expand the body at least: 10% over the settled state of the body, said catalyst being characterized by the ability to promote at least hydrogenation reactions;

said body being characterized by a fairly well-defined upper level of solid particles;

said hydrogen gas .being charged in an amount at least in the range of 700-2500 s.c.f. per barrel of gas oil introduced to obtain the desired degree of conversion and product quality;

said catalytic reaction zone being maintained at a temperature of between about 500 F. and about 900 F. and a pressure in the range of 1000 to 3000 p.s.i.; withdrawing a conversion product mixture from said reaction zone at a point above said upper level of solid particles;

distilling said conversion products mixture to separate the material boiling below about 750 F. from the heavier material; recycling said heavier material to said reaction zone as part of the liquid gas oil charge at a ratio of about 2 to 1 on oil feed and;

continuing said recycle operation until essentially all of the introduced gas oil has been converted to products boiling below about 750 F.;

controlling the combination of the above conditions to maintain a continuous liquid phase in the reactor;

and controlling the reaction rate and amount of recycle oil recirculated so that the temperature difference between the reactor inlet and outlet is less than about 10 F. whereby the product distribution is approximately:

Butane and normally gaseous fractions wt. percent 6 Light naphtha vol. percent 12 Heavy naphtha400 end point vol. percent..- 27 400650 F. fraction vol.percent 2. The process of claim 1 wherein the reaction zone is maintained at a temperature of about 650725 F.

References Cited UNITED STATES PATENTS 3,215,617 11/1965 Burch et a1 208-59 DELBERT E. GANTZ, Primary Examiner. 

